Phosphorous Modified Molecular Sieves, Their Use in Conversion of Organics to Olefins

ABSTRACT

The present invention is a phosphorous modified zeolite (A) made by a process comprising in that order:
         selecting a zeolite with low Si/Al ratio (advantageously lower than 30) among H +  or NH 4   + -form of MFI, MEL, FER, MOR, clinoptilolite, said zeolite having been made preferably without direct addition of organic template;   steaming at a temperature ranging from 400 to 870° C. for 0.01-200h;   leaching with an aqueous acid solution containing the source of P at conditions effective to remove a substantial part of Al from the zeolite and to introduce at least 0.3 wt % of P;   separation of the solid from the liquid;   an optional washing step or an optional drying step or an optional drying step followed by a washing step;
           a calcination step.   
               

     The present invention also relates to a process (hereunder referred as “XTO process”) for making an olefin product from an oxygen-containing, halogenide-containing or sulphur-containing organic feedstock wherein said oxygen-containing, halogenide-containing or sulphur-containing organic feedstock is contacted with the above catalyst (in the XTO reactor) under conditions effective to convert at least a portion of the oxygen-containing, halogenide-containing or sulphur-containing organic feedstock to olefin products (the XTO reactor effluent). 
     The present invention also relates to a process (hereunder referred as “combined XTO and OCP process”) to make light olefins from an oxygen-containing, halogenide-containing or sulphur-containing organic feedstock comprising:
     contacting said oxygen-containing, halogenide-containing or sulphur-containing organic feedstock in the XTO reactor with the above catalyst at conditions effective to convert at least a portion of the feedstock to form an XTO reactor effluent comprising light olefins and a heavy hydrocarbon fraction;   separating said light olefins from said heavy hydrocarbon fraction;   contacting said heavy hydrocarbon fraction in the OCP reactor at conditions effective to convert at least a portion of said heavy hydrocarbon fraction to light olefins.

FIELD OF THE INVENTION

The present invention relates to phosphorus modified molecular sieves aswell as their use in conversion of organics to light olefins. Moreprecisely the P-modified zeolites of the invention are obtained fromcrystalline aluminosilicates having been synthesized preferably withouttemplate. This provides a lower catalyst cost and makes a preparationprocedure more environmentally friendly. The P-modified zeolites of theinvention are useful as catalysts in a variety of processes includingcracking, hydrocracking, isomerization, reforming, dewaxing, alkylation,transalkylation, conversion of oxygenates (or halogenide-containing orsulphur-containing organic compounds) to light olefins.

The limited supply and increasing cost of crude oil has prompted thesearch for alternative processes for producing hydrocarbon products. Onesuch process is the conversion of oxygen-containing (by way of examplemethanol), halogenide-containing or sulphur-containing organic compoundsto hydrocarbons and especially light olefins (by light olefins is meantC₂ to C₄ olefins) or gasoline and aromatics. In the present applicationthe conversion of said oxygen-containing (also referred as oxygenates),halogenide-containing or sulphur-containing organic compounds tohydrocarbons and especially light olefins is referred as XTO process.The interest in the XTO process is based on the fact that feedstock's,especially methanol can be obtained from coal, biomass, hydrocarbonresidues, petcoke, organic waste or natural gas by the production ofsynthesis gas which is then processed to produce methanol. The XTOprocess can be combined with an OCP (olefins cracking process) processto increase production of olefins. The XTO process produces lightolefins such as ethylene and propylene as well as heavy hydrocarbonssuch as butenes and above. These heavy hydrocarbons are cracked in anOCP process to give mainly ethylene and propylene.

BACKGROUND OF THE INVENTION

In accordance with U.S. Pat. No. 3,911,041, methanol or dimethyl etheris subjected to the action, at a temperature of at least about 300° C.,with a catalyst comprising a crystalline aluminosilicate zeolite havinga silica to alumina ratio of at least about 12, a constraint index ofabout 1 to 12, and containing phosphorus incorporated with the crystalstructure thereof in an amount of at least about 0.78 percent by weight.The amount of the phosphorus incorporated with the crystal structure ofthe zeolite may be as high as about 4.5 percent by weight. The zeolite,preferably, also has a dried crystal density of not less than about 1.6grams per cubic centimeter. The crystalline aluminosilicate zeolitehaving a silica to alumina ratio of at least about 12 is first convertedto the hydrogen form, then phosphorus is introduced by reaction with aphosphorus-containing compound having a covalent or ionic constituentcapable of reacting or exchanging with hydrogen ion and thereafterheating. There is no steaming of the zeolite prior to introduction ofphosphorus. Preferably, prior to reacting the zeolite with thephosphorus-containing compound, the zeolite is dried. Drying can beeffected in the presence of air. Elevated temperatures may be employed.

In accordance with U.S. Pat. No. 5,573,990 methanol and/or dimethyletheris converted in presence of a catalyst which contains at least 0.7% byweight of phosphorus and at least 0.97% by weight of rare earth elementsincorporated within the structure of the catalyst. Preferably the amountof phosphorus is comprised between 0.7 and 5% by weight. The phosphoruscontent in the catalyst is most preferably comprised between 1.3 and1.7% by weight. The rare earth elements incorporated with the crystalstructure of the catalyst are preferably rich in lanthanum, the contentof lanthanum in the catalyst being preferably comprised between 2.5 and3.5% by weight. The zeolite ZSM-5 based catalyst presents a mole ratioSiO₂/Al₂O₃ comprised between 40 and 80, a crystal size comprised between1 and 10 μm and adsorption capacities of n-hexane and water 10-11% byweight and 6-7% by weight respectively. Said ZSM-5 is synthesized in thepresence of a template, then is converted to the hydrogen form by ionexchange with hydrochloric acid. The zeolite HZSM-5 prepared asdescribed above is impregnated in aqueous phosphoric acid solution underreduced pressure preferably comprised between 0.08 and 0.09 MPa for 2-3hours. It is dried at <110° C. for 3-5 hours and calcined at about 540°C. for about 3 hours, the phosphorus content of the obtained productPZSM-5 being 0.7-5% (by weight). There is no steaming of the zeoliteprior to introduction of phosphorus. The feedstock methanol comprisessteam in a ratio methanol/steam 10-50/90-50, the examples are made witha ratio 30/70.

U.S. Pat. No. 6,797,851 uses at least two different zeolite catalysts toproduce an olefin composition from an oxygenate, for example, twodifferent ZSM-type catalysts, to produce olefin having a significantquantity of ethylene and propylene. The catalysts can be mixed togetherin one reactor, arranged in separate beds, or used in separate reactorsin series. It is desirable that one of the zeolite catalysts contains aZSM-5 molecular sieve. The ZSM-5 molecular sieve is selected from thegroup consisting of an unmodified ZSM-5, a phosphorous modified ZSM-5, asteam modified ZSM-5 having a micropore volume reduced to not less than50% of that of the unsteamed ZSM-5, and mixtures thereof. It is alsodesirable to have a second zeolite catalyst which contains a zeolitemolecular sieve selected from the group consisting of 10-ring zeolitessuch as ZSM-22, ZSM-23, ZSM-35, ZSM-48, and a mixture thereof. In oneembodiment, the zeolite employed in the first stage of the above processis ZSM-5 having a silica to alumina molar ratio of at least 250, asmeasured prior to any treatment of the zeolite to adjust itsdiffusivity. According to one embodiment, the zeolite is modified with aphosphorous containing compound to control reduction in pore volume.Alternatively, the zeolite is steamed, and the phosphorous compound isadded prior to or after steaming. After contacting with thephosphorus-containing compound, the porous crystalline material,according to one embodiment, is dried and calcined to convert thephosphorus to an oxide form. One or more inert diluents may be presentin the oxygenate feedstock. Preferred diluents are water and nitrogen.Water can be injected in either liquid or vapor form. For example, theprocess may be conducted in the presence of water such that the molarratio water to methanol in the feed is from about 0.01:1 to about 10:1.

US20060106270A1 relates to a process wherein the average propylene cycleselectivity of an oxygenate to propylene (OTP) process using adual-function oxygenate conversion catalyst is substantially enhanced bythe use of a combination of: 1) moving bed reactor technology in thehydrocarbon synthesis portion of the OTP flow scheme in lieu of thefixed bed technology of the prior art; 2) a hydrothermally stabilizedand dual-functional catalyst system comprising a molecular sieve havingdual-function capability dispersed in a phosphorus-modified aluminamatrix containing labile phosphorus and/or aluminum anions; and 3) acatalyst on-stream cycle time of 400 hours or less. The use of a mixtureof a zeolitic catalyst system with a non-zeolitic catalyst system isdescribed. This mixed catalyst embodiment can be accomplished eitherusing a physical mixture of particles containing the zeolitic materialwith particles containing the non-zeolitic material or the catalyst canbe formulated by mixing the two types of material into the phosphorusmodified aluminum matrix in order to form particles having bothingredients present therein. In either case the preferred combination isa mixture of ZSM-5 or ZSM-11 with SAPO-34 in relative amounts such thatZSM-5 or ZSM-11 comprises 30 to 95 wt % of the molecular sieve portionof the mixture with a value of about 50 to 90 wt % being especiallypreferred. It doesn't describe phosphorus modified molecular sieves. Adiluent is preferably used in order to control partial pressure of theoxygenate reactant in the OTP conversion zone and in order to shift theoverall reaction selectivity towards propylene. An especially preferreddiluent for use is steam since it is relatively easily recovered fromthe product effluent stream utilizing condensation techniques. Theamount of diluent used will be selected from the range from about 0.1:1to 5:1 moles of diluent per mole of oxygenate and preferably 0.5:1 to2:1 in order to lower the partial pressure of the oxygenates to a levelwhich favours production of propylene.

EP448000 relates to a process for the conversion of methanol ordimethylether into light olefins in presence of water vapour over asilicoaluminate of the pentasil structure of at least Si/Al ratio of 10,producing at least 5 wt % of ethylene, at least 35 wt % of propylene andat most 30 wt % butenes by (1) using a total pressure of 10 to 90 kPa,(2) a weight ratio of water to methanol of 0.1 to 1.5, (3) a reactortemperature of 280 to 570° C. and (4) a proton-containing catalyst ofthe pentasil-type, having an alkali-content of at most 380 ppm, lessthan 0.1 wt % of ZnO and less than 0.1 wt % of CdO and a BET surfacearea of 300 to 600 m2/gram and a pore volume of 0.3 to 0.8 cm3/gram.

The phosphorus modified molecular sieves of the present invention isprepared based on zeolite with low Si/Al ratio (advantageously below 30)preferably synthesized without direct addition of organic template, thenthe zeolite is subjected to a steam treatment at high temperature beforea leaching step with acid solution containing the source of phosphoruswhich removes advantageously at least 10% of the Al from the zeolite andwhich leads to at least 0.3 wt % of P on the zeolite. It has been foundthat phosphorus acid are very efficient in complexing theextra-framework aluminiumoxides and hence removing them from the zeolitesolid material. Unexpectedly, a larger quantity of phosphorus than whatcould be expected from the typical pore volume of the zeolite andassuming that the pores of the zeolites are filled with the usedphosphorus acid solution, stays in the solid zeolite material. Thechemical functionalities of aluminum with phosphorus in the P-zeoliteinhibit the further dealumination of zeolites, which, in turn, increasestheir stability and selectivity.

The zeolite can be MFI, MOR, MEL, clinoptilolite or FER crystallinealuminosilicate molecular sieves having a low initial Si/Al ratio(advantageously below 30) and preferably synthesized without directaddition of organic directing agent.

The method consists in steaming followed by leaching by a solution ofphosphoric acid or by any acid solution containing the source of P. Itis generally known by the persons in the art that steam treatment ofzeolites, results in aluminium that leaves the zeolite framework andresides as aluminiumoxides in and outside the pores of the zeolite. Thistransformation is known as dealumination of zeolites and this term willbe used throughout the text. The treatment of the steamed zeolite withan acid solution results in dissolution of the extra-frameworkaluminiumoxides. This transformation is known as leaching and this termwill be used throughout the text. Then the zeolite is separated,advantageously by filtration, and optionally washed. A drying step canbe envisaged between filtering and washing steps. The solution after thewashing can be either separated, by way of example, by filtering fromthe solid or evaporated.

The residual P-content is adjusted by P-concentration in the leachingsolution, drying conditions, and washing procedure if any. Thisprocedure leads to dealumination of zeolites and retention of P.Advantageously, at least 0.3 wt % of P is retained after dealuminationon zeolite. Both factors dealumination and the retention of P stabilizethe lattice aluminium in the zeolitic lattice, thus avoiding furtherdealumination. This leads to higher hydrothermal stability, tuning ofmolecular sieves properties and adjustment of acid properties. Thedegree of dealumination can be adjusted by the steaming and leachingconditions.

The P-modified zeolites of this recipe are obtained based on cheapcrystalline alumosilicates with low Si/Al ratio preferably synthesizedwithout direct addition of organic template. This provides a lower finalcatalyst cost and makes a preparation procedure more environmentallyfriendly. The recipe simplifies the procedure for P-ZSM preparation andallows adjusting the Si/Al ratio and P-content in the catalyst. Thecatalysts show high C3-yield, high C3-/C2-ratio, high stability, highC3's purity and reduced selectivity to paraffin's and to aromatic inXTO. These catalysts provide also the additional flexibility forethylene and C4+ recycling for additional propylene production. Theaverage propylene yield can be substantially enhanced by using thesecatalysts in a combination of XTO and OCP process.

BRIEF DESCRIPTION OF THE INVENTION

The present invention is a phosphorous modified zeolite (A) made by aprocess comprising in that order:

-   -   selecting a zeolite with low Si/Al ratio (advantageously lower        than 30) among H⁺ or NH₄ ⁺-form of MFI, MEL, FER, MOR,        clinoptilolite, said zeolite having been made preferably without        direct addition of organic template;    -   steaming at a temperature ranging from 400 to 870° C. for        0.01-200 h;    -   leaching with an aqueous acid solution containing the source of        P at conditions effective to remove a substantial part of Al        from the zeolite and to introduce at least 0.3 wt % of P;    -   separation of the solid from the liquid;    -   an optional washing step or an optional drying step or an        optional drying step followed by a washing step;    -   a calcination step.

The zeolite can be made with the help of seeds techniques but withouttemplate, the seeds could have been made with a template which meansthat the zeolite is made without direct addition of a template.

Advantageously the steaming step and the leaching step are consecutive,there is no intermediate steps such as, by way of example, contact withsilica powder and drying.

Further to the leaching, the separation of the liquid from the solid isadvantageously made by filtering at a temperature between 0-90° C.,centrifugation at a temperature between 0-90° C., evaporation orequivalent.

Optionally, the zeolite can be dried after separation before washing.Advantageously said drying is made at a temperature between 40-600° C.for 1-10 h. This drying can be processed either in a static condition orin a gas flow. Air, nitrogen or any inert gases can be used. Optionallyfurther to the leaching step and the separation the zeolite is dried ata temperature between 40 and 600° C.

The washing step can be performed either during the filtering(separation step) with a portion of cold (<40° C.) or hot water (>40 but<90° C.) or the solid can be subjected to a water solution (1 kg ofsolid/4 liters water solution) and treated under reflux conditions for0.5-10 h followed by evaporation or filtering.

Final calcination step is performed advantageously at the temperature400-700° C. either in a static condition or in a gas flow. Air, nitrogenor any inert gases can be used.

Advantageously (A) when contacted with an oxygen-containing,halogenide-containing or sulphur-containing organic feedstock is capableto make an olefin product.

The present invention also relates to catalyst consisting of the above(A) or comprising the above (A).

The present invention also relates to a process (hereunder referred as“XTO process”) for making an olefin product from an oxygen-containing,halogenide-containing or sulphur-containing organic feedstock whereinsaid oxygen-containing, halogenide-containing or sulphur-containingorganic feedstock is contacted with the above catalyst (in the XTOreactor) under conditions effective to convert at least a portion of theoxygen-containing, halogenide-containing or sulphur-containing organicfeedstock to olefin products (the XTO reactor effluent). It is desirableto have a substantially 100% conversion of the organic compound in theXTO reactor. This conversion rate is adjusted by optimization of contacttime and the frequency of regeneration of the catalyst.

According to a specific embodiment the XTO reactor effluent comprisinglight olefins and a heavy hydrocarbon fraction is sent to afractionation section to separate said light olefins from said heavyhydrocarbon fraction; said heavy hydrocarbon fraction is recycled in theXTO reactor at conditions effective to convert at least a portion ofsaid heavy hydrocarbon fraction to olefin products.

With regards to said effluent of the XTO process, “light olefins” meansethylene and propylene and the “heavy hydrocarbon fraction” is definedherein as the fraction containing hydrocarbons having a molecular weightgreater than propane, which means hydrocarbons having 4 carbon atoms ormore and written as C₄ ⁺.

According to another embodiment of the invention said olefin products(the effluent of the XTO) are fractionated to form a stream comprisingessentially ethylene and at least a part of said stream is recycled inthe XTO reactor to increase the propylene production and then theflexibility of ethylene vs propylene production.

According to another embodiment of the invention both ethylene and theC4+ can be recycled in the XTO reactor.

The present invention also relates to a process (hereunder referred as“combined XTO and OCP process”) to make light olefins from anoxygen-containing, halogenide-containing or sulphur-containing organicfeedstock comprising:

contacting said oxygen-containing, halogenide-containing orsulphur-containing organic feedstock in the XTO reactor with the abovecatalyst at conditions effective to convert at least a portion of thefeedstock to form an XTO reactor effluent comprising light olefins and aheavy hydrocarbon fraction;separating said light olefins from said heavy hydrocarbon fraction;contacting said heavy hydrocarbon fraction in the OCP reactor atconditions effective to convert at least a portion of said heavyhydrocarbon fraction to light olefins. It is desirable to have asubstantially 100% conversion of the organic compound in the XTOreactor. This conversion rate is adjusted by optimization of contacttime and the frequency of regeneration of the catalyst.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates an embodiment with effluent of an XTO reactor passedto a fractionator, and a bottoms of the fractionator sent to an OCPreactor.

FIG. 2 illustrates methane yield and propylene yield vs time on stream(TOS).

DETAILED DESCRIPTION OF THE INVENTION

As regards (A) and the selected zeolite, advantageously it is acrystalline alumosilicate of the MFI family or the MEL family. Anexample of MFI silicates is ZSM-5. An example of an MEL zeolite isZSM-11 which is known in the art. Other examples are described by theInternational Zeolite Association (Atlas of Zeolite Structure Types,1987, Butterworths).

Crystalline silicates are microporous crystalline inorganic polymersbased on a framework of XO₄ tetrahydra linked to each other by sharingof oxygen ions, where X may be trivalent (e.g. Al, B, . . . ) ortetravalent (e.g. Ge, Si, . . . ). The crystal structure of acrystalline silicate is defined by the specific order in which a networkof tetrahedral units are linked together. The size of the crystallinesilicate pore openings is determined by the number of tetrahedral units,or, alternatively, oxygen atoms, required to form the pores and thenature of the cations that are present in the pores. They possess aunique combination of the following properties: high internal surfacearea; uniform pores with one or more discrete sizes; ionexchangeability; good thermal stability; and ability to adsorb organiccompounds. Since the pores of these crystalline alumosilicates aresimilar in size to many organic molecules of practical interest, theycontrol the ingress and egress of reactants and products, resulting inparticular selectivity in catalytic reactions. Crystallinealumosilicates with the MFI structure possess a bi-directionalintersecting pore system with the following pore diameters: a straightchannel along [010]: 0.53-0.56 nm and a sinusoidal channel along [100]:0.51-0.55 nm. Crystalline alumosilicates with the MEL structure possessa bi-directional intersecting straight pore system with straightchannels along [100] having pore diameters of 0.53-0.54 nm.

Advantageously the selected MFI, MEL, FER, MOR, clinoptilolite (or H⁺ orNH₄ ⁺-form MFI, MEL, FER, MOR, clinoptilolite) has an initial atomicratio Si/Al of 30 or lower and preferably ranging from 4 to 30. Theconversion to the H⁺ or NH₄ ⁺-form is known per se and is described inU.S. Pat. No. 3,911,041 and U.S. Pat. No. 5,573,990.

In the steam treatment step, the temperature is preferably from 420 to870° C., more preferably from 480 to 760° C. The pressure is preferablyatmospheric pressure and the water partial pressure may range from 13 to100 kPa. The steam atmosphere preferably contains from 5 to 100 vol %steam with from 0 to 95 vol % of an inert gas, preferably nitrogen. Thesteam treatment is preferably carried out for a period of from 0.05 to200 hours, more preferably from 0.05 to 50 hours. The steam treatmenttends to reduce the amount of tetrahedral aluminium in the crystallinesilicate framework by forming alumina.

The leaching with an aqueous acid solution containing the source of P isadvantageously made under reflux conditions, meaning boiling temperatureof the solution.

Amount of said acid solution is advantageously between 2 and 10 litersper kg of zeolite. A typical leaching period is around 0.5 to 24 hours.Advantageously the aqueous acid solution containing the source of P inthe leaching step has a pH of 3, advantageously 2, or lower.Advantageously said aqueous acid solution is phosphorus acids, a mixtureof phosphorus acids and organic or inorganic acid or mixtures of saltsof phosphorus acids and organic or inorganic acids. The phosphorus acidsor the corresponding salts can be of the phosphate ([PO₄]³⁻, beingtribasic), phosphite ([HPO₃]²⁻, being dibasic), or hypophosphite([H₂PO₂]¹⁻, being monobasic), type. Of the phosphate type also di orpolyphosphates ([P_(n)O_(3n+1)]^((n+2)−)) can be used. The other organicacids may comprise an organic acid such as citric acid, formic acid,oxalic acid, tartaric acid, malonic acid, succinic acid, glutaric acid,adipic acid, maleic acid, phthalic acid, isophthalic acid, fumaric acid,nitrilotriacetic acid, hydroxyethylenediaminetriacetic acid,ethylenediaminetetracetic acid, trichloroacetic acid trifluoroaceticacid or a salt of such an acid (e.g. the sodium salt) or a mixture oftwo or more of such acids or salts. The other inorganic acids maycomprise an inorganic acid such as nitric acid, hydrochloric acid,methansulfuric acid, sulfuric acid or a salt of such an acid (e.g. thesodium or ammonium salts) or a mixture of two or more of such acids orsalts.

Advantageously the final P-content of (A) is at least 0.3 wt % andpreferably between 0.3 and 7 w %. Advantageously at least 10% of Al, inrespect to parent zeolite MFI, MEL, FER, MOR and clinoptilolite, havebeen extracted and removed from the zeolite by the leaching. Theresidual P-content is adjusted by P-concentration in the leachingsolution, drying conditions and a washing procedure if any. A dryingstep can be envisaged between filtering and washing steps.

Then the zeolite either is separated from the washing solution or isdried without separation from the washing solution. Said separation isadvantageously made by filtration. Then the zeolite is calcined, by wayof example, at 400° C. for 2-10 hours.

The solid (A) of the present invention can be used as itself as acatalyst. In another embodiment it can be formulated into a catalyst bycombining with other materials that provide additional hardness orcatalytic activity to the finished catalyst product. Materials which canbe blended with (A) can be various inert or catalytically activematerials, or various binder materials. These materials includecompositions such as kaolin and other clays, various forms of rare earthmetals, phosphates, alumina or alumina sol, titania, zirconia, quartz,silica or silica sol, and mixtures thereof. These components areeffective in densifying the catalyst and increasing the strength of theformulated catalyst. The catalyst may be formulated into pellets,spheres, extruded into other shapes, or formed into a spray-driedparticles. The amount of (A) which is contained in the final catalystproduct ranges from 10 to 90 weight percent of the total catalyst,preferably 20 to 70 weight percent of the total catalyst.

With regards to the XTO process, the catalyst of the invention isparticularly suited for the catalytic conversion of oxygen-containing,halogenide-containing or sulphur-containing organic compounds tohydrocarbons. Accordingly, the present invention also relates to amethod for making an olefin product from an oxygen-containing,halogenide-containing or sulphur-containing organic feedstock whereinsaid oxygen-containing, halogenide-containing or sulphur-containingorganic feedstock is contacted with the above catalyst under conditionseffective to convert the oxygen-containing, halogenide-containing orsulphur-containing organic feedstock to olefin products (the effluent ofthe XTO). Said effluent comprises light olefins and a heavy hydrocarbonfraction.

In this process a feedstock containing an oxygen-containing,halogenide-containing or sulphur-containing organic compound contactsthe above described catalyst in a reaction zone of a reactor atconditions effective to produce light olefins, particularly ethylene andpropylene. Typically, the oxygen-containing, halogenide-containing orsulphur-containing organic feedstock is contacted with the catalyst whenthe oxygen-containing, halogenide-containing or sulphur-containingorganic compounds is in vapour phase. Alternately, the process may becarried out in a liquid or a mixed vapour/liquid phase. In this process,converting oxygen-containing, halogenide-containing orsulphur-containing organic compounds, olefins can generally be producedat a wide range of temperatures. An effective operating temperaturerange can be from about 200° C. to 700° C. At the lower end of thetemperature range, the formation of the desired olefin products maybecome markedly slow. At the upper end of the temperature range, theprocess may not form an optimum amount of product. An operatingtemperature of at least 300° C., and up to 600° C. is preferred.

The pressure also may vary over a wide range. Preferred pressures are inthe range of about 5 kPa to about 5 MPa, with the most preferred rangebeing of from about 50 kPa to about 0.5 MPa. The foregoing pressuresrefer to the partial pressure of the oxygen-containing,halogenide-containing, sulphur-containing organic compounds and/ormixtures thereof.

The process can be carried out in any system using a variety oftransport beds, although a fixed bed or moving bed system could be used.Advantageously a fluidized bed is used. It is particularly desirable tooperate the reaction process at high space velocities. The process canbe conducted in a single reaction zone or a number of reaction zonesarranged in series or in parallel. Any standard commercial scale reactorsystem can be used, for example fixed bed, fluidised or moving bedsystems. After a certain time on—stream the catalyst needs to beregenerated. This regeneration can be carried out in a separate reactoror in the same reactor. In case of a moving bed or fluidised bedreactor, a part of the catalyst is continuously or intermittentlywithdrawn from the conversion reactor and sent to a second reactor forregeneration. After the regeneration, the regenerated catalyst iscontinuously or intermittently sent back to the conversion reactor. Incase of fixed bed reactor the reactor is taken off-line forregeneration. Generally this requires a second spare reactor that cantake over the conversion into light olefins. After regeneration thefixed bed reactor is in stand-by until the spare reactor needsregeneration and the regenerated reactor takes over the conversion.Regeneration is carried out by injecting an oxygen-containing streamover the catalyst at sufficient high temperature to burn the depositedcoke on the catalyst. The commercial scale reactor systems can beoperated at a weight hourly space velocity (WHSV) of from 0.1 hr⁻¹ to1000 hr⁻¹.

One or more inert diluents may be present in the feedstock, for example,in an amount of from 1 to 95 molar percent, based on the total number ofmoles of all feed and diluent components fed to the reaction zone.Typical diluents include, but are not necessarily limited to helium,argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen, water,paraffins, alkanes (especially methane, ethane, and propane), aromaticcompounds, and mixtures thereof. The preferred diluents are water andnitrogen. Water can be injected in either liquid or vapour form.

According to a specific embodiment essentially no water (or steam) isinjected as diluent of the feedstock sent to the XTO reactor. However itmeans that the feedstock can contain the water already contained in thefresh oxygen-containing, halogenide-containing or sulphur-containingorganic feedstock or the steam used to engage proper flowing of catalystin fluidised bed of moving bed reactors of the XTO reactor.

The oxygenate feedstock is any feedstock containing a molecule or anychemical having at least an oxygen atom and capable, in the presence ofthe above catalyst, to be converted to olefin products. The oxygenatefeedstock comprises at least one organic compound which contains atleast one oxygen atom, such as aliphatic alcohols, ethers, carbonylcompounds (aldehydes, ketones, carboxylic acids, carbonates, esters andthe like). Representative oxygenates include but are not necessarilylimited to lower straight and branched chain aliphatic alcohols andtheir unsaturated counterparts. Examples of suitable oxygenate compoundsinclude, but are not limited to: methanol; ethanol; n-propanol;isopropanol; C₄-C₂₀ alcohols; methyl ethyl ether; dimethyl ether;diethyl ether; di-isopropyl ether; formaldehyde; dimethyl carbonate;dimethyl ketone; acetic acid; and mixtures thereof. Representativeoxygenates include lower straight chain or branched aliphatic alcohols,their unsaturated counterparts. Analogously to these oxygenates,compounds containing sulphur or halides may be used. Examples ofsuitable compounds include methyl mercaptan; dimethyl sulfide; ethylmercaptan; di-ethyl sulfide; ethyl monochloride; methyl monochloride,methyl dichloride, n-alkyl halides, n-alkyl sulfides having n-alkylgroups of comprising the range of from about 1 to about 10 carbon atoms;and mixtures thereof. Preferred oxygenate compounds are methanol,dimethyl ether, or a mixture thereof.

In XTO effluent among the olefins having 4 carbon atoms or more thereare greater than 50 weight % of butenes. More than 80% by weight andadvantageously more than 85% of the hydrocarbons having 4 carbon atomsor more are C4 to C8 olefins.

According to a specific embodiment the XTO reactor effluent comprisinglight olefins and a heavy hydrocarbon fraction is sent to afractionation section to separate said light olefins from said heavyhydrocarbon fraction; said heavy hydrocarbon fraction is recycled in theXTO reactor at conditions effective to convert at least a portion ofsaid heavy hydrocarbon fraction to olefin products.

With regards to said effluent of the XTO process, “light olefins” meansethylene and propylene and the “heavy hydrocarbon fraction” is definedherein as the fraction containing hydrocarbons having a molecular weightgreater than propane, which means hydrocarbons having 4 carbon atoms ormore and written as C₄ ⁺.

According to another embodiment of the invention said olefin products(the effluent of the XTO) are fractionnated to form a stream comprisingessentially ethylene and at least a part of said stream is recycled inthe XTO reactor to increase the propylene production and then theflexibility of ethylene vs propylene production. Advantageously theratio of ethylene to the oxygen-containing, halogenide-containing orsulphur-containing organic feedstock is 1.8 or less.

According to another embodiment of the invention both ethylene and theC4+ can be recycled in the XTO reactor.

The present invention also relates to a process (hereunder referred as“combined XTO and OCP process”) to make light olefins from anoxygen-containing, halogenide-containing or sulphur-containing organicfeedstock comprising:

contacting said oxygen-containing, halogenide-containing orsulphur-containing organic feedstock in the XTO reactor with the abovecatalyst at conditions effective to convert at least a portion of thefeedstock to form an XTO reactor effluent comprising light olefins and aheavy hydrocarbon fraction;separating said light olefins from said heavy hydrocarbon fraction;contacting said heavy hydrocarbon fraction in the OCP reactor atconditions effective to convert at least a portion of said heavyhydrocarbon fraction to light olefins.

The effluent of the XTO reactor comprising light olefins and a heavyhydrocarbon fraction is sent to a fractionation section to separate saidlight olefins from said heavy hydrocarbon fraction. With regards to saideffluent of the XTO process, “light olefins” means ethylene andpropylene and the “heavy hydrocarbon fraction” is defined herein as thefraction containing hydrocarbons having a molecular weight greater thanpropane, which means hydrocarbons having 4 carbon atoms or more andwritten as C₄ ⁺. It is desirable to have a substantially 100% conversionof the organic compound in the primary reactor. This conversion rate isadjusted by optimization the contact time and the frequency of theregeneration of the catalyst.

With regards to the OCP process, said process is known per se. It hasbeen described in EP 1036133, EP 1035915, EP 1036134, EP 1036135, EP1036136, EP 1036138, EP 1036137, EP 1036139, EP 1194502, EP 1190015, EP1194500 and EP 1363983 the content of which are incorporated in thepresent invention. The heavy hydrocarbon fraction produced in the XTOreactor is converted in the OCP reactor, also called an “olefin crackingreactor” herein, to produce additional amounts of ethylene andpropylene.

According to a specific embodiment the catalysts found to produce thisconversion comprise a crystalline alumosilicate of the MFI family or theMEL family. These alumosilicates have been described above in thedescription of (A).

According to a specific embodiment of the invention the catalysts foundto produce this heavy conversion are the catalysts consisting of theabove (A) or comprising the above (A). They can be the same as thecatalysts of the XTO reactor or although they are in the description of(A) they can be different of the XTO catalyst because of the startingzeolite, the P content etc . . . .

According to another embodiment the catalysts found to produce thisheavy conversion are disclosed in the patents EP 1036133, EP 1035915, EP1036134, EP 1036135, EP 1036136, EP 1036138, EP 1036137, EP 1036139, EP1194502, EP 1190015, EP 1194500 and EP 1363983.

The crystalline alumosilicate catalyst has structural and chemicalproperties and is employed under particular reaction conditions wherebythe catalytic cracking of the C₄ ⁺ olefins readily proceeds. Differentreaction pathways can occur on the catalyst. Under the processconditions, having an inlet temperature of around 400° to 600° C.,preferably from 520° to 600° C., yet more preferably 540° to 580° C.,and an olefin partial pressure of from 0.1 to 2 bars, most preferablyaround atmospheric pressure. Olefin catalytic cracking may be understoodto comprise a process yielding shorter molecules via bond breakage.

The MFI catalyst having a high silicon/aluminum atomic ratio for use inthe OCP reactor of the present invention may be manufactured by removingaluminum from a commercially available crystalline silicate. A typicalcommercially available silicalite has a silicon/aluminum atomic ratio ofaround 120. The commercially available MFI crystalline silicate may bemodified by a steaming process which reduces the tetrahedral aluminum inthe crystalline silicate framework and converts the aluminum atoms intooctahedral aluminum in the form of amorphous alumina. Although in thesteaming step aluminum atoms are chemically removed from the crystallinesilicate framework structure to form alumina particles, those particlescause partial obstruction of the pores or channels in the framework.This inhibits the olefinic cracking processes of the present invention.Accordingly, following the steaming step, the crystalline silicate issubjected to an extraction step wherein amorphous alumina is removedfrom the pores and the micropore volume is, at least partially,recovered. The physical removal, by a leaching step, of the amorphousalumina from the pores by the formation of a water-soluble aluminumcomplex yields the overall effect of de-alumination of the MFIcrystalline silicate. In this way by removing aluminum from the MFIcrystalline silicate framework and then removing alumina formed therefrom the pores, the process aims at achieving a substantiallyhomogeneous de-alumination throughout the whole pore surfaces of thecatalyst. This reduces the acidity of the catalyst and thereby reducesthe occurrence of hydrogen transfer reactions in the cracking process.The reduction of acidity ideally occurs substantially homogeneouslythroughout the pores defined in the crystalline silicate framework. Thisis because in the olefin-cracking process hydrocarbon species can enterdeeply into the pores. Accordingly, the reduction of acidity and thusthe reduction in hydrogen transfer reactions which would reduce thestability of the MFI catalyst are pursued throughout the whole porestructure in the framework. The framework silicon/aluminum ratio may beincreased by this process to a value of at least about 180, preferablyfrom about 180 to 1000, more preferably at least 200, yet morepreferably at least 300 and most preferably around 480.

The MEL or MFI crystalline silicate catalyst may be mixed with a binder,preferably an inorganic binder, and shaped to a desired shape, e.g.extruded pellets. The binder is selected so as to be resistant to thetemperature and other conditions employed in the catalyst manufacturingprocess and in the subsequent catalytic cracking process for theolefins. The binder is an inorganic material selected from clays,silica, metal oxides such as ZrO₂ and/or metals, or gels includingmixtures of silica and metal oxides. The binder is preferablyalumina-free, although aluminum in certain chemical compounds as inAlPa_(t)'s may be used as the latter are quite inert and not acidic innature. If the binder which is used in conjunction with the crystallinesilicate is itself catalytically active, this may alter the conversionand/or the selectivity of the catalyst. Inactive materials for thebinder may suitably serve as diluents to control the amount ofconversion so that products can be obtained economically and orderlywithout employing other means for controlling the reaction rate. It isdesirable to provide a catalyst having a good crush strength. This isbecause in commercial use, it is desirable to prevent the catalyst frombreaking down into powder-like materials. Such clay or oxide bindershave been employed normally only for the purpose of improving the crushstrength of the catalyst. A particularly preferred binder for thecatalyst of the present invention comprises silica or AlPO₄.

The relative proportions of the finely divided crystalline silicatematerial and the inorganic oxide matrix of the binder can vary widely.Typically, the binder content ranges from 5 to 95% by weight, moretypically from 20 to 50% by weight, based on the weight of the compositecatalyst. Such a mixture of crystalline silicate and an inorganic oxidebinder is referred to as a formulated crystalline silicate. In mixingthe catalyst with a binder, the catalyst may be formulated into pellets,spheres, extruded into other shapes, or formed into a spray-driedpowder.

In the catalytic cracking process of the OCP reactor, the processconditions are selected in order to provide high selectivity towardspropylene or ethylene, as desired, a stable olefin conversion over time,and a stable olefinic product distribution in the effluent. Suchobjectives are favoured by the use of a low acid density in the catalyst(i.e. a high Si/Al atomic ratio) in conjunction with a low pressure, ahigh inlet temperature and a short contact time, all of which processparameters are interrelated and provide an overall cumulative effect.

The process conditions are selected to disfavour hydrogen transferreactions leading to the formation of paraffins, aromatics and cokeprecursors.

The process operating conditions thus employ a high space velocity, alow pressure and a high reaction temperature. The LHSV ranges from 0.5to 30 hr⁻¹, preferably from 1 to 30 hr⁻¹. The olefin partial pressureranges from 0.1 to 2 bars, preferably from 0.5 to 1.5 bars (absolutepressures referred to herein). A particularly preferred olefin partialpressure is atmospheric pressure (i.e. 1 bar). The heavy hydrocarbonfraction feedstock is preferably fed at a total inlet pressuresufficient to convey the feedstocks through the reactor. Said feedstockmay be fed undiluted or diluted in an inert gas, e.g. nitrogen or steam.Preferably, the total absolute pressure in the second reactor rangesfrom 0.5 to 10 bars. The use of a low olefin partial pressure, forexample atmospheric pressure, tends to lower the incidence of hydrogentransfer reactions in the cracking process, which in turn reduces thepotential for coke formation which tends to reduce catalyst stability.The cracking of the olefins is preferably performed at an inlettemperature of the feedstock of from 400° to 650° C., more preferablyfrom 450° to 600° C., yet more preferably from 540° C. to 590° C.,typically around 560° to 585° C.

In order to maximize the amount of ethylene and propylene and tominimize the production of methane, aromatics and coke, it is desired tominimize the presence of diolefins in the feed. Diolefin conversion tomonoolefin hydrocarbons may be accomplished with a conventionalselective hydrogenation process such as disclosed in U.S. Pat. No.4,695,560 hereby incorporated by reference.

The OCP reactor can be a fixed bed reactor, a moving bed reactor or afluidized bed reactor. A typical fluid bed reactor is one of the FCCtype used for fluidized-bed catalytic cracking in the oil refinery. Atypical moving bed reactor is of the continuous catalytic reformingtype. As described above, the process may be performed continuouslyusing a pair of parallel “swing” reactors. The heavy hydrocarbonfraction cracking process is endothermic; therefore, the reactor shouldbe adapted to supply heat as necessary to maintain a suitable reactiontemperature. Online or periodic regeneration of the catalyst may beprovided by any suitable means known in the art.

The various preferred catalysts of the OCP reactor have been found toexhibit high stability, in particular being capable of giving a stablepropylene yield over several days, e.g. up to ten days. This enables theolefin cracking process to be performed continuously in two parallel“swing” reactors wherein when one reactor is operating, the otherreactor is undergoing catalyst regeneration. The catalyst can beregenerated several times.

The OCP reactor effluent comprises methane, light olefins andhydrocarbons having 4 carbon atoms or more. Advantageously said OCPreactor effluent is sent to a fractionator and the light olefins arerecovered. Advantageously the hydrocarbons having 4 carbon atoms or moreare recycled at the inlet of the OCP reactor, optionally mixed with theheavy hydrocarbon recovered from the effluent of the XTO reactor.Advantageously, before recycling said hydrocarbons having 4 carbon atomsor more at the inlet of the OCP reactor, said hydrocarbons having 4carbon atoms or more are sent to a second fractionator to purge theheavies. In a preferred embodiment the light olefins recovered from theeffluent of the XTO reactor and the light olefins recovered from thefractionator following the OCP reactor are treated in a common recoverysection.

Optionally, in order to adjust the propylene to ethylene ratio of thewhole process (XTO+OCP), ethylene in whole or in part can be recycledover the OCP reactor and advantageously converted into more propylene.This ethylene can either come from the fractionation section of the XTOreactor or from the fractionation section of the OCP reactor or fromboth the fractionation section of the XTO reactor and the fractionationsection of the OCP reactor or even from the optional common recoverysection.

Optionally, in order to adjust the propylene to ethylene ratio of thewhole process (XTO+OCP), ethylene in whole or in part can be recycledover the XTO reactor where it combines with the oxygen-containing,halogenide-containing or sulphur-containing organic feedstock to formmore propylene. This ethylene can either come from the fractionationsection of the XTO reactor or from the fractionation section of the OCPreactor or from both the fractionation section of the XTO reactor andthe fraction section of the OCP reactor or even from the optional commonrecovery section.

These ways of operation allow to respond with the same equipment andcatalyst to market propylene to ethylene demand.

FIG. 1 illustrates a specific embodiment of the invention. The effluentof the XTO reactor is passed to a fractionator 11. The overhead, a C1-C3fraction including the light olefins is sent via line 2 to a commonrecovery section (not shown). The bottoms (the heavy hydrocarbonfraction) are sent via line 3 to the OCP reactor. The effluent of theOCP reactor is sent via line 10 to a fractionator 8. The overhead, aC1-C3 fraction including the light olefins, is sent via line 9 to acommon recovery section (not shown). The bottoms, hydrocarbons having 4carbon atoms or more, are sent to a fractionator 5. The overhead,hydrocarbons having 4 to substantially 5 carbon atoms are recycled vialine 4 at the inlet of the OCP reactor. The bottoms, hydrocarbons havingsubstantially 6 carbon atoms or more, are purged via line 6.

The method of making the olefin products from an oxygenate feedstock caninclude the additional step of making the oxygenate feedstock fromhydrocarbons such as oil, coal, tar sand, shale, biomass and naturalgas. Methods for making oxygenate feedstocks are known in the art. Thesemethods include fermentation to alcohol or ether, making synthesis gas,then converting the synthesis gas to alcohol or ether. Synthesis gas canbe produced by known processes such as steam reforming, autothermalreforming and partial oxidization in case of gas feedstocks or byreforming or gasification using oxygen and steam in case of solid (coal,organic waste) or liquid feedstocks. Methanol, methylsulfide andmethylhalides can be produced by oxidation of methane with the help ofdioxygen, sulphur or halides in the corresponding oxygen-containing,halogenide-containing or sulphur-containing organic compound.

One skilled in the art will also appreciate that the olefin productsmade by the oxygenate-to-olefin conversion reaction using the molecularsieve of the present invention can be polymerized optionally with one ormore comonomers to form polyolefins, particularly polyethylenes andpolypropylenes. The present invention relates also to said polyethylenesand polypropylenes.

EXAMPLES Example 1

A sample of zeolite ZSM-5 with Si/Al=13 with a crystal size <1 μm inH-form synthesized without template has been obtained from TRICAT®. Thesample is hereinafter identified as Comparative I.

Example 2

A sample of zeolite ZSM-5 described in example 1 was steamed at 550° C.for 48 h. Steamed solid was treated by 3.14M solution of H₃PO4 for 18 hunder reflux condition (4.2 liter/1 kg of zeolite). Then the solid wasseparated by filtering from the solution. Obtained solid was dried at110° C. for 16 h and calcined at 400° C. for 10 h.

(Atomic ratio Si/Al—25, P-content 5.6 wt %). The sample is hereinafteridentified as Sample A.

Example 3

This example contains a procedure for preparation of dealuminated ZSM-5by steaming and leaching with phosphorus acid. However almost all P wasremoved by washing. Thus the zeolite was dealuminated by the same waybut contained a very low P amount.

A sample was prepared using the same procedure as in example 2, exceptthe sample was washed 6×2000 ml of distilled water per kg of zeolite.The sample was dried right away after the filtering at 110° C. for 16 hand calcined at 400° C. for 10 h.

(Atomic ratio Si/Al—26, P-content 0.2 wt %). The sample is hereinafteridentified as Comparative II.

Example 4

This example contains a procedure for P-ZSM-5 preparation with aP-content close to in sample A. However in this case the order ofsteaming and treatment with phosphorus acid was inversed. In the sametime, this recipe produces non-leached sample modified with phosphorous,meaning a total Al content in the zeolite after this treatment remainedunchanged.

100 g of zeolite ZSM-5 described in example 1 was impregnated with 23 gof H₃PO₄ (85 wt % in water) at room temperature and dried at 110° C. for16 h. Then the sample was calcined at 400° C. for 10 h. Finally,obtained solid was steamed at 550° C. for 48 h. (Atomic ratio Si/Al is13 and P content 4 w %).

The sample is hereinafter identified as Comparative III.

Example 5

This example presents non-dealuminated ZSM-5 with the same atomic Si/Alratio and P-content as in Sample A, but the P was introduced byimpregnation.

A sample of zeolite ZSM-5 with Si/Al=25 with a crystal size <1 μm inH-form has been obtained from Zeolyst International. The sample wascalcined at 550° C. for 6 h. Then the zeolite was impregnated by theincipient wetness method with an aqueous solution of (NH₄)₂HPO₄ with atarget to introduce 5% of P in the sample. 60 g of the calcined zeolitewas impregnated with a solution containing 47 g water and 13,461 g(NH₄)₂HPO₄. Finally the P-zeolite was dried overnight at 110° C. andcalcined at 600° C. for 10 h. (Atomic ratio Si/Al is 25 and P content 5w %).

The sample is hereinafter identified as Comparative VI.

Example 6

In this example P was introduced in zeolite during the leaching(dealumination).

A sample of zeolite ZSM-5 with Si/Al=13.8 with a crystal size <1 μm inH-form has been obtained from Zeolyst. The sample was steamed at 680° C.for 2 h. The steamed solid was treated by 3.14M solution of H₃PO4 for 18h under reflux condition (4.2 liter/1 kg of zeolite). Then the solid wasseparated by filtering from the solution and dried at 400° C. for 3 h inair. Then the dried sample was subjected in a contact with a hot watersolution under reflux condition for 2 h. Then the solid was separated byfiltering from the solution and dried right away at 110° C. for 16 h andsteamed at 600° C. for 2 h (Atomic ratio Si/Al—21, P-content 1.8 wt %).

The sample is hereinafter identified as Sample B.

Example 7

This example presents a recipe for P-ZSM-5 dealuminated by steaming andtreatment with phosphorous acid and some additional P was introducedafter dealumination by impregnation.

A sample was prepared using the same procedure as in example 6, exceptafter the leaching and filtering steps the solid was washed with 2000 mlof distilled water per kg of zeolite (P content >0.3 wt %).

Then 53.67 g of the sample was mixed with 225.4 ml of water contained4,064 g of (NH₄)H₂PO₄. Then the solution containing the zeolite and(NH₄)H₂PO₄ was evaporated under stirring. Obtained solid was dried at110° C. for 16 h and steamed at 600° C. for 2 h. (Atomic ratio Si/Al—21,P content 2.10 wt %).

The sample is hereinafter identified as Sample C.

Example 8-9

XTO Conditions (XTO in the Table)

Catalyst tests were performed on 2 g catalyst samples with a puremethanol feed in a fixed-bed, down flow stainless-steel reactor.Catalyst powders was pressed into wafers and crushed to 35-45 meshparticles. Prior to catalytic run all catalysts were heated in flowingN₂ (5 Nl/h) up to the reaction temperature. Analysis of the products hasbeen performed on-line by a gas chromatograph equipped with a capillarycolumn. The catalyst performances were compared at substantially fullmethanol conversion, under equal conditions and maximum propylene yield.The results are displayed on carbon and water free basis.

OCP Conditions (XTO+OCP in the Table)

The feedstock which contains substantially non cyclic olefins C4+ (theheavy hydrocarbon fraction) was subjected to catalytic cracking (thesecond reactor) in the presence of an aluminosilicate catalyst in afixed bed reactor at 575° C., LHSV=10 h⁻¹, P=1.5 bara. This catalystcomprises a commercially available silicalite which had been subjectedto a dealumination treatment by combination of steaming with acidtreatment so as provide Si/Al ratio ˜250. A detailed procedure ofcatalyst preparation is described in above cited EP1194502 B1.

The OCP performance has been simulated using a mathematic modelemploying conversion factors deduced from numerous testing of differentfeedstocks. Based on the stream composition going to the OCP reactor andon the required purges an optimum stream of C4 and heaviers are recycledaround the OCP reactor. The lines under “OCP feed non cyclic olefinsC4+” display the heavy hydrocarbon flow rate sent to the OCP (the secondreactor). The lines under “XTO+OCP” display the ethylene and propyleneproduced by the combination of the primary reactor (MTO) and the secondreactor (OCP).

The values in table are the weight percent on carbon basis.

Example 8 Comparative I Comparative II Sample A ZSM-5 P-ZSM-5 P-ZSM-5Si/Al 13 26 25 P, % 0 <0.2 5.6 XTO T, ° C. 450 450 450 WHSV, h⁻¹ 1.6 1.61.6 P, barg 0.5 0.5 0.5 C1 (methane) 3.4 1.2 1.3 Paraffins 51.6 37.011.2 Olefins 29.2 54.0 69.2 Dienes 0.0 0.3 1.5 Aromatics 19.2 8.8 6.6C3-/C2- 1.6 2.3 4.8 C2- + C3- 6.7 23 28 ethylene 2.6 7 5 propylene 4.116 23 OCP feed (non cyclic olefins C4+) Σ olefins 18 28 40 XTO + OCPC3-/C2- 2.8 2.9 4.3 C2- + C3- 20 43 58 ethylene 5 11 11 propylene 15 3247

Example 9 Comparative III Comparative VI Sample A P-ZSM-5 P-ZSM-5P-ZSM-5 Si/Al 13 25 25 P, % 4 5 5.6 XTO T, ° C. 550 550 550 WHSV, h⁻¹1.6 1.6 1.6 P, barg 0.5 0.5 0.5 C1 (methane) 6.4 2.3 1.6 Paraffins 11.74.7 5.5 Olefins 73.0 82.1 86.1 Dienes 1.1 1.6 1.9 Aromatics 14.2 6.0 5.2C3-/C2- 2.0 4.8 5.0 C2- + C3- 47 45 46 ethylene 15 8 8 propylene 32 3738 OCP feed (Non cyclic olefins C4+) Σ olefins 25 36 39 XTO + OCPC3-/C2- 2.5 4.5 4.4 C2- + C3- 66 71 75 ethylene 19 13 14 propylene 47 5861Methane yield and propylene yield vs time on stream (TOS) are on FIG. 2.It is clear from the FIG. 2 that the comparative examples 3 and 4 do notresult in stable propylene production.

Example 10

XTO Conditions (XTO in the Table)

Catalyst tests were performed on 2 g catalyst samples with a puremethanol feed in a fixed-bed, down flow stainless-steel reactor.Catalyst powders was pressed into wafers and crushed to 35-45 meshparticles. Prior to catalytic run all catalysts were heated in flowingN₂ (5 Nl/h) up to the reaction temperature. Analysis of the products hasbeen performed on-line by a gas chromatograph equipped with a capillarycolumn. The catalyst performances were compared at full methanolconversion, under equal conditions and maximum propylene yield. Theresults are displayed on carbon and water free basis.

Example 10 Sample C Sample B P-ZSM-5 P-ZSM-5 Si/Al 21 21 P, % 2.1 1.8 T,° C. 550 550 WHSV, h⁻¹ 1.6 1.6 P, barg 0.5 0.5 XTO C1 (methane) 3.4 2.0Paraffins 5.9 5.6 Olefins 83.7 80.6 Dienes 1.3 0.9 Aromatics 8.3 11.8C3-/C2- 4.5 2.7 C2- + C3- 48 52 ethylene 9 14 propylene 39 38 OCP feed(Non cyclic olefins C4+) Σ olefins 35 28

Example 11 (OCP)

Catalyst tests were performed on 10 ml (˜6 g) of catalyst grains (35-45meshes) (Sample A) loaded in the tubular reactor. The feedstock whichcontains substantially non cyclic olefins C4 (˜60%) was subjected tocatalytic cracking in the presence of catalyst in a fixed bed reactor at550° C., LHSV=2 h⁻¹, P=1.5 bara. The results are in table 1 hereunder.The values in table 1 are the weight percent on carbon basis.

The data given below illustrate a cracking activity of the Sample A inC4 olefins conversion to propylene and ethylene at the same temperatureand pressure as in XTO reactor.

SAMPLE A feed effluent Paraffins 41.1 41.5 Olefins 58.8 55.5 Dienes 0.00.7 Aromatics 0.0 2.3 C1 (methane) 0.0 0.4 Ethylene 0.0 5.0 Propane 0.61.4 Propylene 0.3 20.8 Butenes 57.4 19.2

1-54. (canceled)
 55. A process comprising: contacting anhalogenide-containing feedstock in a first reactor with a catalyst underconditions effective to convert the halogenide-containing feedstock to afirst reactor effluent comprising olefin products, wherein the catalystcomprises a phosphorus modified zeolite having a P content rangingbetween 0.3 and 7 weight percent, wherein the phosphorus modifiedzeolite is made by a process comprising: selecting a zeolite having aSi:Al atomic ratio of 30 or less, wherein the zeolite is selected fromthe group consisting of MFI, MEL, FER, MOR, and clinoptilolite, andwherein the selected zeolite is in the H+ or NH4+ form; steaming thezeolite at a temperature ranging from 400° C. to 870° C. for 0.01 h to200 h; leaching the zeolite with an aqueous acid solution containing asource of P at conditions effective to remove a substantial amount of Alfrom the zeolite and to introduce more than 0.3 wt % of P, wherein theaqueous acid solution containing the source of P comprises a salt of aphosphate ([PO4]³⁻), an acid or a corresponding salt of a phosphite([HPO₃]²⁻), an acid or a corresponding salt of a hypophosphite([H₂PO₂]¹⁻), an acid or a corresponding salt of a diphosphate, or anacid or a corresponding salt of a polyphosphate; separating the zeolitefrom the aqueous acid solution; and calcining the zeolite; wherein thefirst reactor effluent comprises light olefins and a heavy hydrocarbonfraction and is sent to a first fractionator to separate the lightolefins from the heavy hydrocarbon fraction, and wherein the heavyhydrocarbon fraction is sent to a second reactor at conditions effectiveto convert at least a portion of the heavy hydrocarbon fraction to lightolefins, wherein a catalyst in the second reactor is the same as thecatalyst in the first reactor.
 56. The process of claim 55, wherein thefirst reactor effluent comprises light olefins and a heavy hydrocarbonfraction and is sent to a first fractionator to separate the lightolefins from the heavy hydrocarbon fraction, and wherein the heavyhydrocarbon fraction is recycled to the first reactor at conditionseffective to convert at least a portion of the heavy hydrocarbonfraction to olefin products.
 57. The process of claim 55, wherein theolefin products include ethylene and propylene that are fractionated toform a stream comprising ethylene, and wherein at least a part of thestream comprising ethylene is recycled to the first reactor to increasepropylene production.
 58. The process of claim 55, wherein a secondreactor effluent is sent to a second fractionator and the light olefinsare recovered, and wherein heavy hydrocarbons having 4 or more carbonatoms are recycled to the second reactor and mixed with the heavyhydrocarbons recovered from the first reactor effluent.
 59. The processof claim 58, wherein the heavy hydrocarbons having 4 or more carbonatoms are sent to a third fractionator to remove a heavy hydrocarbonstream comprising C6+ hydrocarbons prior to recycling to the secondreactor.
 60. The process of claim 58, wherein the olefin productsinclude ethylene and propylene, wherein ethylene is recycled to thesecond reactor to adjust a propylene to ethylene production ratio, andwherein the ethylene is recycled from the first fractionator, the secondfractionator, both the first fractionator and the second fractionator,or a common recovery section.
 61. The process of claim 58, wherein theolefin products include ethylene and propylene, wherein ethylene isrecycled to the first reactor to adjust a propylene to ethyleneproduction, and wherein the ethylene is recycled from the firstfractionator, the second fractionator, both the first fractionator andthe second fractionator, or from a common recovery section.
 62. Theprocess of claim 55, wherein the olefin products include ethylene, andwherein the ethylene is polymerized with one or more comonomers.
 63. Theprocess of claim 55, wherein the olefin products include propylene, andwherein the propylene is polymerized with one or more comonomers. 64.The process of claim 63, wherein a ratio of ethylene to theoxygen-containing feedstock fed to the first reactor is 1.8 or less. 65.The process of claim 55, wherein the first reactor is operated at atemperature ranging from about 200° C. to 700° C., and wherein a partialpressure of the oxygen-containing feedstock ranges from about 5 kPa toabout 5 MPa.
 66. The process of claim 55, wherein thehalogenide-containing feedstock is contacted in the first reactor withthe catalyst in the presence of an inert diluent, wherein the inertdiluent is present in an amount ranging from 1 to 95 molar percent basedon a total number of moles of the inert diluent and thehalogenide-containing feedstock.
 67. The process of claim 55, whereingreater than 50 weight % of olefins having 4 carbon atoms or more in thefirst reactor effluent are butenes.
 68. The process of claim 55, whereinmore than 80% by weight of hydrocarbons having 4 carbon atoms or more inthe first reactor effluent are C4 to C8 olefins.
 69. The process ofclaim 55, wherein the aqueous acid solution containing the source of Pcomprises phosphorous acid (H₃PO₃).
 70. The process of claim 55, whereinthe phosphorus modified zeolite comprises ZSM-5.